Hydroforming with platinum-on-alumina catalyst



J. B. MALLOY 2,916,433

HYDROFORMING WITH PLATINUM-ON-ALUMINA CATALYST Dec. 8, 1959 3 Sheets-Sheet 1 Filed Jan. 31, 1955 AGZINVJOUdEIU Ql a auosav RI r '0 aoiavsa 1N VEN TOR. John B. Mal/0y Dec. 8, 1959 J. B. MALLOY 2,916,433

HYDROFORMING WITH PLATINUM-ON-ALUMINA CATALYST Filed Jan. 31, 1955 3 Sheets-Sheet 2 OCH/YE NUMBER OF BOTTOMS A/VD OVER/15405 FROM 93.6 OOTANE' ULTRAFORMATE (M0. Nap/film} OFR-R I 80 GUMULA TIVE OVERHEAD 400 VOL. 2 DIST/LL50 TRUE BOILING PO/NT ULTRAFORMATE 200 DIST/LLAT/ON I l l 0 I00 VOL. 9; DIST/LL50 INVEN TOR. Flg- 2 John a. Mal/0y AT OR/VE Y Dec. 8, 1959 J. B. MALLOY 2,916,433

HYDROFORMING wrru PLATINUM-ON-ALUMINA CATALYST Filed Jan. 31, 1955 s Sheets-Sheet :5

OGTA/VE NUMBER OF BOTTOM? AND OVERHEADS FROM 85.6 OGTA/VE ULTRAFORMATE {M 6'. Nflpbl/Id} "i L f 1 I OCT/WE w. 80

GFR-R CUMULATIVE 70 OVERHEAD I l I l 25 50 75 I00 400 VOL. 2 DIST/LL50 TRUE BOILING POINT zoo UL TRAFORMATE "F DIST/LLATION VOL. 2 DIST/LL50 IN VEN TOR. F 3 John a. Mal/0y ATTOR E United States Patent HYDROFORMING WITH PLATINUM-0N- ALUMINA CATALYST John B. Malloy, Lansing, 111., assignor to Standard Oil Company, Chicago, 111., a corporation of Indiana Application January 31, 1955, Serial No. 485,246 3 Claims. (Cl. 208-65) This invention relates to an improved: method ofv hydroforrning naphthas with platinum-on-alumina catalyst and it pertains more particularly to an improvement in a multistage adiabatic system for hydroforming naphthas of relatively wide boiling range.

Many catalytic processes are known for hydroforming naphthas with platinum-on-alurnina catalysts including Platforming, Ultraforming, Cat-forming, Houdriforming, etc. Platforming is a non-regenerative process operating at pressures of 500 psi. or more and it depends largely on hydrocracking for obtaining product octane numbers upwards of 90 CPR-R. Ultraforming operates at lower pressure and is a regenerative process in which hydrocracking is minimized and the achievement of high octane number is attained largely by dehydrocyclizing paraffi'ns. While my invention may be employed in Platforming, it is particularly applicable to Ultra'forming and to other regenerative platinum-on-alumina hydroforming processes employing catalysts and conditions for maximizing d'ehydrocyclizing and minimizing hydrocracking of the paraffins contained in the naphtha charge.

An object of the invention is to maximize dehydrocyclizing of parafiins in a platinum-on-alumina hydroforming process while minimizing both hydrocracking and coke formation. In other words, the object is to obtain increased yields of higher octane number products from a given naphtha charge and with a given amount of catalyst without sacrificing run lengths. A further object is to enable the use in a platinum-onalumina process of naphthas of higher end points than have heretofore been feasible particularly in non-regen erative processes. An important object is to increase the effectiveness and efficiency of platinum-on-alumina hydroforming and to obtain naphtha products having octane numbers of 95 to 100 or more with a minimum of capital investment and operating expense. Other objects will be apparent as the detailed description of the invention proceeds.

The principal reactions effected by a hydroforming process are (1) formation of aromatics by naphthene dehydrogenation, (2) isomerization of paraflins and naphthenes, (3) formation of aromatics from parafiins by dehydrocyclizing and (4) hydrocracking. The dehydrogenation and isomerization reactions are most rapid so that by the time the naphtha reaches the finalcontacting stage, most of the naphthenes originally contained therein have been converted to aromatics, the product at this point being largely a mixture of paraifinic and aromatic hydrocarbons. To improve octane number further, the concentration of aromatics in the liquid must be increased by dehydrocyclizing paraflins or by hydrocracking paraffins to lighter components and gases. Operation at pressures substantially below 500 p.s.i. and

ice

preferably not higher than about 400 psi. favors the desired dehydrocyclizing reaction.

The equilibrium ratio of aromatic to parafiins increases with temperature and increases even to agreater extent as the size of the parafiin increases as shown by the following table:

Equilibrium Ratio of Aromatic t0 Paraffin at 300 p.s.i.g., 5,000

It will thus be apparent that equilibrium will prevent high conversion of low-boiling paratfins unless high reforming temperatures are employed. Use of high reforming temperatures at low pressure has heretofore been undesirable because of excessive deposition of coke on the catalyst, for which the higher-boiling hydrocarbons are largely responsible. In accordance with this invention, the high-boiling aromatics formed by dehydrocyclizati'on in the initial stage or stages of the hydroforrning process are separated from intermediate product so that only the low-boiling and incompletely dehydrocyclized parafiins are subjected to the final hydroforming stage. The presence of the low boiling components F. to 240 F. boiling range) are desirable in the initial stage not only because they are partially converted therein without appreciably contributing to catalyst deactivation or coke formation but because they serve as heat carriers for supplying the necessary endothermic heat of dehydrocyclizing. By removing the high boiling fraction from the intermediate product prior to the final stage, I avoid hydrocracking of previously formed high boiling products, condensation or degradation of initially formed high boiling products, and the deactivation of catalyst in the final stage which would otherwise be caused by excessive coke deposits, and I provide in the final stage a higher proportion of hydrogen to naphtha which is undergoing conversion in order to further suppress coke formation and in order to provide additional heat carrier so that maximum dehydrocyclizing of hexanes and heptanes is obtainable. By blending the final light. product from the last hydroforming stage with heavy product separated from the intermediate stage, I obtain a total product of considerably higher octane number Without appreciable yield loss due to hydrocracking and without excessive catalyst deactivation or carbon formation.

For minimizing investment and operating cost the eflluent from an initial or intermediate contacting stage is partially condensed to enable separation of hydrogen and hydrocarbons boiling below about 240 F. as a separate stream from initial condensate. This initial condensate must then be fractionated to remove overhead most of the hydrocarbons having a true boiling point below about 240 F. and in some cases to remove hydrocarbons. boiling up to 300 F. The gas stream and the overhead from the intermediate fractionator are then combined and reheated to the temperature of about 900 to 980 F. and contacted with platinum alumina catalyst in the final contacting zone at a liquid hourly space velocv mediate fractionation. From the standpoint of obtaining ity in the range of about .1 to in order to obtain the desired dehydrocyclizing of the lighter paraflins.

While the amount of high boiling product removed in the intermediate stage will depend to a certain extent on the nature of the charging stock, the catalyst and the octane level to be achieved, intermediate product upwards of about 300 F. true boiling point should in all cases be removed and it may be desirable to remove products having a true boiling point above 240 F. With the usual naphtha charges this means that about 70 to 30 percent of the product naphtha will be separated as bottoms in the intermediate fractionation step and only the remaining 30 to 70 percent will be contacted in the final hydroforming stage. t

The invention will be more clearly understood from the following detailed description of a specific example read in conjunction with the accompanying drawings which form a part of this specification and in which:

Figure 1 is a schematic flow diagram of an improved Ultraformer system embodying my invention,

Figure 2 is a chart showing the octane number of bottoms and overhead fractions from 93.6 octane number Ultraformate, and

Figure 3 is a chart showing the octane number of bottoms and overhead fractions from 85.6 octane number Ultraformate. In the lower part of Figure 2, I have shown a true boiling point curve A for a representative 93.6 CFR-R octane number Ultraformate obtained by conventional Ultraforming of an M-C naphtha boiling chiefly in the range of about 150 to 380 F. In the upper part of Figure 2, I have shown as curve B the octane number of cumulative overhead fractions of said Ultraformate starting with the first 10 percent of overhead products. In curve C I have shown the octane number of bottoms remaining after the removal of the indicated volume percent of overhead material. Straight line D is the octane number of the whole Ultraformate.

It will be observed that the initial 30 or 40 percent of the overhead product was characterized by a very low octane number, the fraction from about 10 to 30 percent of the cumulative overhead having an octane number of the order of about 80 CFR-R, indicating a low conversion to aromatics. The bottoms fraction, after removal of this 30 percent of initial light ends, has an octane number of about 99. If 50 percent of the product is taken overhead, the 50 percent bottoms has an octane number above 100. In this case about 70 percent or at least about 50 percent of the product should be separated as bottoms and only the lighter 30 to 50 percent subjected to a further hydroforming step.

Figure 3 is similar to Figure 2 except that it shows the true boiling point distillation curve B of an 85.6 octane number Ultraformate which is obtainable by 'Ultraforming under less severe conditions or with a fewer number of stages. Here again it will be observed from the cumulative overhead curve F that the first 30 percent is characterized by a very low octane number and that it would be desirable in any event to remove this light fraction and subject it to a further hydroforming step. If it is desirable to obtain 100 octane number product, in this case it may be desirable to remove about 70 percent of the low boiling constituents, i.e. substantially all material boiling above 300 F., and to subject this 70 percent of cumulative overhead to further hydroforming in one or more steps in order to obtain more complete dehydrocyclizing of the paraffins contained therein. Bottoms, after removal of components having a true boiling point below 300 R, will have at octane number upwards of '100 as shown at the right end of curve G as compared with a total Ultraformate octane number of 85.6 as shown by straight line H.

From the standpoint of coke formation, it is most important that the components of the intermediate charge .boiling above about 300. F. be separated by the intermaximum octane numbers, a fraction boiling below 240 F. should receive more severe or extensive hydroforming than hydrocarbons having true boiling points above 240 F.

The invention will be described as applied to the Ultraforming of an M-C naphtha boiling in the range of about 150 to 380 F. containing about .03 weight percent sulfur having a CFR-R octane number of about 46 and having a composition roughly of about 40 percent naphthenes, 8 percent aromatics and 52 percent paraflins, all by volume. Such naphtha charge is introduced by line 10 together with about 2,000 to 8,000, e.g. 4,000, cubic feet per barrel of a recycled hydrogen stream from line 11, is heated in coil 12 to a transfer line temperature in the range of about 900 to 950 F. under a pressure of about 320 p.s.i. and introduced through transfer line 13 into number 1 reactor 14. For a 9,000 barrels per day unit reactor 14 may contain about 10.5 tons or 420 cubic feet of platinum-on-alumina catalyst in the form of inch pellets. Such catalyst may be prepared by contacting an aqueous solution of chloroplatinic acid containing 3.5 grams of platinum per liter with an ammonium sulfide solubilizing agent and combining the resulting true or colloidal solution with hydrous alumina prepared as taught in US. Reissue 22,196, the relative amounts of the two components being such as to give a final catalyst containing about .1 to 1, e.g. about .6, percent platinum by weight on a dry basis. If desired the alumina may be aged in an aqueous ammonia solution at a pH of about 10 to 11 for a period of several days to a week and then dried and impregnated with the platinum chloride in the presence or absence of a small amount of added aluminum chloride, the resulting catalyst in any event being dried and calcined. The alumina may contain a small amount of halogen such as fluorine or chlorine but it should be substantially free from sodium, iron and molybdenum. Since platinum-on-alumina catalysts and their methods of preparation are known to those skilled in the art, they will not be described in further detail.

The effiuent stream from the number 1 reactor is discharged by line 15 to reheater 16 wherein it is heated to a temperature of the order of 900 to 950 F. and introduced by line 17 into number 2 reactor 18 which contains the same amount of the same type of catalyst employed in reactor number 1. Here, due to pressure drop in the system, the pressure may be about 300 p.s.i. and usually the temperature drop through this reactor is somewhat less than in the number 1 reactor.

The efiiuent from the number 2 reactor is withdrawn by line 19 through heat exchanger 20 and cooler 21 to separator 22 wherein hydrogen and a portion of the hydrocarbons boiling below about 240 F. are separated and removed through line 23. The condensate from separator 22 is passed by line 24 through heat exchanger 20 to intermediate fractionator 25 which is provided with usual reboiler 26 and which is designed to separate from the intermediate condensate most and preferably all of the hydrocarbons having true boiling points below about 240 F. In some cases it is desirable to also take overhead the hydrocarbons having true boiling points up to about 300 F., particularly when the portion in the range of 240 to 300 F. may be materially improved in octane number by further hydroforming. Thus in the case of an intermediate fraction having about 85 octane number when the desired product is to have an octane number upwards of 95, it is preferred to take overhead about percent of the intermediate product, i.e. substantially all hydrocarbons having a true boiling point below about 300 F. A part of the overhead from intermediate fractionator 25 is passed through cooler 27 to receiver 28 and recycled by pump 29 for use as reflux.

The net overhead is introduced by line 30 together with the hydrogcn fraction from line 23 to coil 31 of the final reheating furnace wherein the mixture is heated to a temperature of about 900 to 980 F., e.g. 950 F., and introduced by line 32 into the number 3 reactor 33. This reactor likewise contains the same amount of the same type of catalyst as is employed in reactors 1 and 2 and, due to pressure drop, the reaction pressure in this case will be somewhat below 300 p.s.i.g. The final reactor efiluent is withdrawn through line 34, heat exchanger 35 and cooler 36 to separator 37 which is preferably operated at about 100 F. or lower and from which separated hydrogen is removed through line 38. The net hydrogen produced may be withdrawn through line 39 to absorber 40 and condensables may be removed from the net hydrogen stream and finally removed from the system by line 41 by introducing absorber oil through line 42 and withdrawing enriched absorber oil through line 43. The recycle hydrogen stream is returned by compressor 44 through exchanger 35 in line 11 for admixture with the incoming naphtha charge.

The final condensate from separator 37 is introduced by line 45 to depropanizer 46 from which C and lighter hydrocarbons are withdrawn overhead through line 47 and condenser 48 to receiver 49, a part of the condensate from the receiver being returned by pump 50 and line 51 for use as reflux and the remainder withdrawn from the system through line 51a. The stabilized light product withdrawn from the base of the depropanizer through line 52 is admixed with the intermediate product withdrawn from the base of intermediate fractionator 25 through line 53 so that the final product withdrawn through line 54 is a naphtha of full boiling range having an octane number of at least about 95 CFRR.

In the low pressure Ultraforming system a swing reactor 55 is preferably employed which reactor contains the same amount of the same type of catalyst as reactors number 1, 2 and 3. The swing reactor is provided with a charging stock inlet line 56 which may be selectively connected to lines 14a, 17a or 32a. It is also provided with an effiuent line 57 which may be selectively connected to lines 15a, 19a or 34a. Thus the swing reactor may be connected in parallel with any one of the other reactors when no catalyst in the system requires regeneration or rejuvenation. When the catalyst in any one of the onstream reactors declines in activity or selectivity to the point that regeneration and rejuvenation is required, it may be replaced by the swing reactor and the deactivated catalyst may be regenerated and rejuvenated by system 58 which is operatively connected by manifolds 59 and 60 to each of the reactors in the system. A preferred method of efiecting regeneration and rejuvenation of catalyst in such a swing reactor system is described and claimed in application Serial No. 416,072, filed March 15, 1954, now US. 2,892,770, and, since no novelty is claimed in this aspect of the operating procedure, it will not be herein described in any further detail. It should be pointed out, however, that by removing the heaviest portion of the intermediate product prior to the final contacting stage, cracking and coke deposition in this final stage is markedly decreased so that the catalyst in the final reactor can remain on-stream for a much longer period of time without necessity of regeneration or rejuvenation even when charging stocks of end points up to 400 or 425 F. are employed.

While a particular example of the process has been described in considerable detail, it will be understood that a single reactor or more than two reactors may be employed prior to intermediate product fractionation and that more than one reactor stage may be employed subsequent to the intermediate product fractionation. The final reactor may be smaller or may contain less catalyst than the initial reactors; while uniformly sized reactors are employed in the example, the invention is applicable to systems wherein reactor sizes are not uniform. Alternative arrangements and operating conditions will be ap- 6 parent from the above description to those skilled in the art.

I claim:

1'. In a regenerative platinum catalyst naphtha hydroforming process wherein a charge stream is heated to about 900 to 950 F. and passed through a first conversion zone, the effluent from the first conversion zone is reheated to a temperature of about 900 to 950 F. and passed through a second conversion zone and a substantial part of the effluent from the second conversion zone reheated to a temperature in the range of about 900 to 980 F. and passed through a final conversion zone, the effluent from the final conversion zone being passed through a cooling zone to a separating zone for separation therefrom of a hydrogen stream which is recycled through a heating zone to the first conversion zone and wherein selected conversion zones are at intervals replaced in any desired sequence with a swing conversion zone which may be selectively connected in parallel with and in place of any of the first, intermediate or final conversion zones so that the catalyst in the latter may be regenerated without interruption of on-stream flow in the defined sequence through the respective zones, the improved method of operation which comprises cooling effluent from the intermediate conversion zone to an extent sufiicient to condense hydrocarbons contained therein which boil above about 300 F., separating the resulting condensate from uncondensed hydrocarbons and hydrogen, fractionating the condensate to segregate from bottoms an overhead stream comprising most of the hydrocarbons having a true boiling point below 240 F., charging to the final reheating and conversion zone the mixture of said overhead stream with hydrogen and uncondensed hydrocarbons from the separating step and blending the bottoms from the fractionating step with product produced in the final conversion zone.

2. In a process for hydroforming naphtha with platinum-on-alumina catalyst wherein a mixture of preheated naphtha vapors and recycled hydrogen is contacted with platinum-on-alumina catalyst under hydroforming conditions in an initial contacting zone, total efiiuent from the initial contacting zone is reheated and contacted with platinum-on-alumina catalyst under hydroforming conditions in an intermediate contacting zone, a mixture of hydrogen and naphtha vapors obtained from efiluent leaving the intermediate contacting zone is heated to a temperature in the range of 900 to 980 F. and contacted with a platinum-on-alumina catalyst under hydroforming conditions in a final contacting zone, a part of the hydrogen separated from the product leaving the final contacting zone being recycled to provide said recycled hydrogen, the improved method of operation which comprises condensing as initial condensate from the eflluent from the intermediate contacting zone hydrocarbons boiling above about 300 F., separating as a gas stream from the initial condensate a mixture of hydrogen and light hydrocarbons boiling in the range of about to 240 F., fractionating the initial condensate to obtain a bottoms fraction amounting to at least about 30 percent of the total hydrocarbons in the stream leaving the intermediate contacting zone and an overhead stream of hydrocarbons having a true boiling range below about 300 F. and combining the gas stream with the overhead stream to form the mixture of hydrogen and naphtha vapors which is heated and contacted with platinum-on-alumina catalyst in the final contacting zone.

3. A multistage adiabatic process for hydroforming naphtha with platinum-on-alumina catalyst wherein there is reheating between stages and wherein the final stage is at a higher average temperature than the initial and intermediate stages, the improved method of operation which comprises removing from the intermediate stage eflluent about 30 to 70 percent of the partially converted naphtha as distillation bottoms prior to the final reheating and contacting stage and reheating and contacting in the final '7 stage only the remainder of the partially converted naphtha together with substantially all of the hydrogen from the intermediate stage whereby dehydrocyclizing of naphthacomponeuts boiling below 300 F. can be increased without degradation of high boiling alkyl aromatic 5 products previously formed.

References Cited in the file of this patent UNITED STATES PATENTS 2,304,183 Layng et al Dec. 8, 1942 8 Mattoz; May 9, 1944 Munday et a1 Jan. 16, 1945 Weikart; June 14, 1955 Hemrninger Mar. 6, 1956 Haensel'et a1. Apr; 3, 1956 Arundale et al. Aug. 7, 1956 Wolf et a1. Dec. 4, 1956 Snuggs et a1. Dec. 4, 1956 

1. IN A REGENERATIVE PLATINUM CATALYST NAPPTHA HYDROFORMING PROCESS WHEREIN A CHARGE STREAM IS HEATED TO ABOUT 900 TO 950* F. AND PASSED THROUGH A FIRST CONVERSION ZONE, THE EFFLUENT FROM THE FIRST CONVERSION ZONE IS REHEATED TO A TEMPERATURE OF ABOUT 900 TO 950* F. AND PASSED THROUGH A SECOND CONVERSION ZONE AND A SUBSTANTIAL PART OF THE EFFUENT FROM THE SECOND CONVERSION ZONE IS REHEATED TO A TEMPERATURE IN THE RANGE OF ABOUT 900 TO 980* F. AND PASSED THROUGH A FINAL CONVERSION ZONE, THE EFFUENT FROM THE FINAL CONVERSION ZONE BEING PASSED THROUGH A COOLING ZONE TO A SEPARATING ZONE FOR SEPARATION THEREFROM OF A HYDROGEN STREAM WHICH IS RECYCLED THROUGH A HEATING ZONE TO THE FIRST CONVERSION ZONE AND WHEREIN SELECTED CONVERSION ZONES ARE AT INTERVALS REPLACED IN ANY DESIRED SEQUENCE WITH A SWING CONVERSION ZONE WHICH MAY BE SELECTIVELY CONNECTED IN PARALLEL WITH AND IN PLACE OF ANY OF THE FIRST, INTERMEDIATE OR FINAL CONVER- 